Thermal conversion of hydrocarbon gases



J. w. THRdCKMoR-roN 2,321.1,251 i THERMAL CONVERSION OF HYDROCARBON GASES May 25, 1943.

Filed Dec. 2B, 1940 BNN.

NBN.

- ATTORNEY.

Patented May 25, 1943 THERMAL CONVERSION OF HYDROCARBON GASES John W. Throckmorton, Wilton, Conn., assignor to The Pure Oil Company, Chicago, Ill., a corporation of Ohio Application December 28, 1940, Serial No. 372,145

4 Claims.

This invention relates to the conversion of hydrocarbon gases to liquid hydrocarbons and is more particularly concerned .with the pyrolytic conversion of hydrocarbon gases under high temperatu-res and superatmospheric pressures.

In the pyrolytic conversion of hydrocarbon gases to gasoline and other higher boiling hydrocarbons, it is common practice to eitherrquench the reaction products leaving the reaction zone by direct contact of the hot reaction products with cooler liquid medium or to cool the reaction products by indirect heat interchange. Both methods of cooling have vtheir disadvantages. Where direct quenching .is practiced prior to indirect heat exchange utilization of the heat in the reaction products for heating the charge to the reaction zone is not as elcient due to the lower mean temperature difference between the two, thereby making the cost of operation greater than it would otherwise be.v

On the other hand, Where indirect cooling is resorted to, as, for example, by passing the charging stock into indirect heat interchange with the reaction products, excess carbon depositionv takes place in the. heat exchanger and transfer line, making necessary more frequent shutdown of the plant, with resultant loss ofl time and efciency. i

It has been discovered that the disadvantages of both methods of cooling reaction products can be to a large extent obviated andthe benefits of both methods realized by carefully 4regulating the extent to which the reaction products are cooled by indirect heat 'exchange prior to quenching the partially cool-ed products With cooling liquid. `It. has been observed that partial `reduction in the temperature of the reaction products can be ob'- tained by indire-ct heatinterchange without any substantial carbon deposition if cooling is vnot carried below the dew point of the `reaction products. By dew point is meant that 1point at which droplets of liquid form and separate out of `the conditions existing.

` reaction mixture under the particular pressurevv VBy utilizing the heat remainingin the hot reacY conditions existing after the reaction products are chilled.

In the pyrolytic conversion of gases to high boiling hydrocarbons, a certain amount of very high boiling material of the nature of tar and asphalt forms. If the reaction products are permitted to cool sufliciently to permit these high boiling materials to condense and drop out of the reaction stream, they deposit in the transfer line or in theheat exchanger, gradually building up out prior to quenching, this material dissolves in or is carried along in suspension in the quench liquid into the fractionating zone and can be eliminated in the liquid body which collects in the zone without causing clogging. p

i It is the primary object of this invention to utilize the heat remaining in the reaction products from a gas conversion zone Without causing undesirable deposition of carbon or carbonaceous material from the reaction products.

It is a further object of this invention to utilize the heatv in the reaction products from a gas conversion zone to preheat the charge to the conversion zone and at the same time prevent undesirable deposition of carbon or carbonaceous mate-f rial from the reaction products.

' Further objects of the invention will become apparent from the following description taken in conjunction with the accompanying drawing of which the single figure is a diagrammatic side` elevational View of apparatus for practicing the invention.

` Referring to the drawing, numeral I indicates aline throughyvhich fresh gas may be charged to the rsystem by means of compressor or other suitable device. From the line l gas passes through valve 3 into line 5 -irom which it enters the bottom of absorber "l, The fresh gas charged to the system may be all or a portion of natural gas, gases produced 'in the pyrolytic and/or catalytic crackng of mineral oil, or a mixture of such gases. The gas passes upwardly through the absorber 1 in countercurrent contact with a descending stream of absorption menstruurn such as light gas oil which enters the upper part of the absorber through the line 9. rThe absorber may be maintained under a superatmospheric pressure ranging from approximately 150 to 400 lbs. per square inch. It will be understood that a series of absorbers may be used or separate absorbers may be used for the fresh gas and the recycle gas and these absorbers operated under different conditions of pressure and/or temperature. These features are all Well known in the art and form no part of the invention.

In the absorber substantially all Cs and higher boiling hydrocarbons are absorbed together with a considerable portion of Cz hydrocarbons. The greater part of the hydrogen, methane and a portion of the C2 hydrocarbons remains unabsorbed.

The gas which is unabsorbed is eliminated from the top of the absorber through line |I controlled by valve I3. The rich absorption menstruum is withdrawn from the bottom of the absorber through line |5 and charged into the stripper I1 where it is denuded of the absorbed gases. The lean absorption menstruum is withdrawn from the stripper through line |9 by means of pump 2| and returned to the top of the absorber through line 9. Make-up oil may be pumped into the absorber when necessary through line 23 controlled by valve 25.

The gases liberated in the stripper which consist primarily of C3 and C4 hydrocarbons with a substantialproportion of C2 hydrocarbons and very little methane and hydrogen, are withdrawn through line 21, passed through cooler 29 where the temperature may be lowered to approximatelyl '15 F., andl then charged to feed tank 33. The stripper I1 is operated under superatmospheric pressure so that the major portion of the gases liberated therein after cooling will condense to liquid and be collected as a liquid in the feed tank 33. The feed tank may be maintained under a pressure ranging from approximately 125 to 375 lbs. per square inch. Any gas which does not condense in the tank 33 is withdrawn therefrom through line 35 controlled by valve 31 and returned by means of pump 38 and line 5 to the absorber.

The liquefied gas is withdrawn from tank 33 through line 39 and a portion thereof charged by means of pump 4| through line 43, heat exchangers 45, 41 and 49, into the inlet of the heating coil 5| In passing through the series of heat exchangers, the charging stock may be heated to a temperature of 450-600 F., depending on the temperature of the reaction products and the extent of heat exchange contact therewith. In the pump 4|, the pressure on the charge may be raised to approximately 500 to 2000 lbs. per square inch. charging stock may be heated to a temperature of approximately 900 to 1200 F., depending on the nature of the charging stock and the nature and yield of the final products desired. Charging stocks of higher olen concentration generally require lower reaction temperatures than do lower olefin containing gases. Likewise, higher boiling parafiinic lhydrocarbons such as butane do not require as high temperatures for conversion as Ydo lower boiling constituentssuch as propane..

From the heating coil `5| the products may pass to a reaction coil 53 of suicient length and cross-sectional area to give the heated products time to Yreact to thedesired extent. The coil 53 is preferably o f greater cross-sectional area than the heating coil 5|, thereby decreasing the velocity of Ythe reaction productsV and increasing the time during which the products remain in the In the heating coil 5| the reaction zone. The reaction coil 53 is preferably placed in a zone which can be heated and/or cooled as necessary to maintain the desired reaction temperature. The amount of heating or cooling necessary is dependent on the concentration of olens in the charging stock. The temperature that is preferably maintained in the reaction coil 53 is approximately the same temperature at which the gases leave the heating coil 5|, or slightly lower. As the reaction products emerge from the reaction coil 53 they pass through the tubes in heat exchanger 49 in indirect heat exchange with the fresh charge. In the heat exchanger 49 the reaction products may be cooled to a temperature of approximately 400 to 800 F., depending upon their composition and the pressure under which they are maintained. Higher pressure and higher concentration of high boiling constituents necessitate higher temperatures to avoid cooling below the dew point. Immediately upon emerging from the heat exchanger 49 the partially cooled reaction products are quenched'by injecting cooler liquid hydrocarbons directly into the stream at point 55. At point 55 the temperature of the reaction products is suddenly reduced from approx- I benefit of the heat in the reaction products in preheating the charge can be realized.

Although any inert liquid may be used as the quenching medium, a substantial portion of which remains liquid under the temperature and pressure .conditions at which the reaction products are quenched, it is preferred to use condensate formed in the process. To this end con'-V densate is withdrawn from the bottom of the flash tower 51 through lines 6| and'63 by means of pump 65. Since this condensate maybe at a temperature of approximately 250500 F., it is rst passed through heat exchangers 41 and 45 in indirect heat. exchange with the charging stock in order to partially coolthe condensate and give up a'portion of` its heat tov thefresh charge 'and is then, if necessary, passed through cooler 61 toVv further reduce its temperature before being injected into the reaction products at the point 55.

The quenching oil or condensate may have'itsA temperature reduced to approximately %225?l F. before contactingit with the reaction Vprodists. to be quenched. A portion of the condensate withdrawn from the bottom of the flash tower 514 may also be charged by pump throughrlne f S9 and cooler 1| intothe upperportion ofthe ash tower 51 to act as reflux liquid.

v A portion of the condensate Withdrawn from the ash tower 51 through line 6| is passed through line 13, controlled by` valve 15, andline 'I1 into the mid-portion of main fractionatingVV tower or stabilizer 19. A portion of the charging stock from feed tank 33 may Y be withdrawn through line 89 by means of pump 8| and charged to the upper portion of tower 19 as reflux liquid# The pressure in flash tower 51 will be some1, what lower than the pressure in the'reacton coil 'I'he quenched able the' taking overhead of the .major portion Y of the vnormally .gaseous constituents and a' portion Jof the gasoline boiling hydrocarbons. The

vapors :and gases which are not condensed in the ash tower 51 'are removed from the upper portion thereof through the line 82 andwa'tery cooler 83 where they are cooled suiiiciently to produce a 4condensate containing Ygasoline 'boiling hydrocarbons, the major portion of the C3-C4 hydrocarbons, a portion of the C2 hydrocarbons and some methane and then ipass to accumulator 85. In` the cooler B3 the vapors may be vcooled to a temperature of approximately 75-100 F. Any gas 4which remains, uncondensed is Withdrawn yfrom the accumulator 35 through line `lili controlled by valve r81 and 'returned through line -5 to the absorber. A portion of the condensate` in the accumulator 85 may -f be returned by means of pump 88 and line 89 to the upper portion of the flash tower 51 as reux liquid. The remainder of the liquid from accumulator 85 is withdrawn through line 9|, controlled by Valve 93, and joins the liquid in line I. 13 from the bottom of; flash tower 51 and enters r stabilizer 19. The pressure is reduced at the valve top temperature of 145 F. and bottom tempera- L ture of 420 F. The condensate in the bottom of the fractionator 19 is heated by being passed into reboiler |2| through line |23. Vapors from the reboiler re-enter the fractionating tower through line |24. The liquid from the reboiler which is composed of gasoline and higher boiling hydrocarbons is withdrawn through line to valve |26 where the pressure is reduced and charged into the fractionator |21.

are taken overhead through line |29, condensed in condenser |3| and passed to accumulator |33. From the accumulator |33 the gasoline distillate is withdrawn by means of pump |35 to storage or to further treatment if required. A portion of the gasoline distillate may be recycled through line |31 to the upper part of fractionator |21 as reflux condensate. The liquid which accumulates in the bottom of the stripper is charged to a, reboiler |39 through line |4| and vapors are returned from the boiler to the fractionator through line |43. The unvaporized residue comprising those hydrocarbons boiling above the gasoline boiling rangey are withdrawn from the reboiler |39 through line |45 by means of pump |41 to storage.

In the actual operation of a plant in accordano with the invention, the feed tank 33 was maintained under a pressure of 235 lbs. per square In the fracn 'tionator |21 the gasoline boiling hydrocarbons inch and at a temperature of F. The composite liquefied gas in feed tank 33 had the following'oomposition in .mole percent:

Hasi-co2 CH4 1.5

C2H4 2.8 02H6 14.4

03H6 22.6 C3Hs 44.2 04H8 '7.3 C4H1o 6.0 Heavier than C4 hydrocarbons 0.9

The liquefied gas was taken from the tank by the pump 4I and raised to a pressure of 700 lbs. per square inch. The temperature of the gas leaving the heat exchanger 49 was 470 F. and vthis tem'- perature was raised to 1080 F. in the heating coil 5| and maintained at this temperature in the reaction coil 53. The pressure at the outlet of the' reaction coil was 530 lbs. per square inch. Reaction products leaving the reaction coil had the following composition in mole percent:

H2s+cog 4 C4H10 Heavier than C4 hydrocarbons 5.1

The temperature of the reaction products was reduced from 1080 F. to 800 F. in the heat exchanger 49 and was quenched from 800 to 400 F. by means of the cool condensate at 55. The condensate used for quench was cooled to 175 F'. The flash tower 51 was maintained under a pressure of 425 lbs. per square inch. The bottom of the tower was maintained at a temperature of 370 F. and the top of the tower was maintained at a temperature of 215 F. The accumulator was maintained under a pressure of 415 lbs. per square inch and a temperature of 85 F. 'I'he ractionating tower 19 was maintained under a pressure of 245 lbs. per square inch with the bottom maintained at a temperature of 420 F. and the top at a temperature of F. The fractionator |21 was operated at a pressure slightly above atmospheric. Under these conditions the unit was able to operate for a period of 25 days without requiring shutdown for cleaning. At the end of this period the deposition in the transfer line from the reaction tubes to the flash tower and in the heat exchanger tubes 49 was found to consist of a little light coke so small in amount as to cause substantially no building up of back pressure.

The exit gas from the absorber had the following composition in mol. percent:

HzS-I-COz 1.5

r om 64.5

C2H4 9.9 C21-I5 24.1

. C3 and heavier hydrocarbons Trace gases intoliquid hydrocarbons at elevated tem-` peratures between approximately 900 and 1200 F. at which the reaction products remain in the c vapor phase, the steps of reducing the reaction products by indirect heat exchange with charging gases to a temperature close to but not below the dew point thereof and then suddenly reducing the temperature of the partially cooled reaction products below the dew point by direct contact with cooler liquid of such nature that a large portion thereof remains liquid at the temperature to which the reaction products are reduced.

2. A process in accordance with claim 1 in which the reaction products are cooled from a temperature above 900 F. to a temperature not substantially below 700 F. by indirect heat inter-'- change and from a temperature not substantially below 700 F. to a temperature substantially below 700 F. by direct contact with cooler liquid of such nature that a large portion thereof remains liquid at the temperature to which the reaction products are-reduced. f-

3. A process in accordance with claimJ 1 in which the reaction products are cooled from. a temperature above 900 F. to a temperature of approximately 800"V F. by indirect heat nterchange and from a temperature of approximately 800' F. to a temperature of approximately 400 F. by direct contact with cooler liquid of such'nature that a large portion thereof remains liquid at the temperature to which the reaction products are reduced.

4. A process in accordance with claim 1 inv which the gases are converted under a pressureabove 500 lbs. per square inch, the reaction products are partially cooled by heat interchange withV charging gas to a temperature of approximately 650 to 800 F. and the partially cooled reaction products quickly cooled by means of direct contact with cooler liquid to a temperature of approximately 250 to 500 F.

JOHN W. IHROCKMORTON 

